Apparatus comprising a catalytic distillation zone comprising a reaction zone with distribution of hydrogen

ABSTRACT

For the treatment of a feed comprising a major portion of olefinic hydrocarbons containing 4 carbon atoms per molecule, including isobutene, 1-butene and 2-butenes in a ratio which substantially corresponds to the thermodynamic equilibrium, in which the feed is treated in a distillation zone comprising a stripping zone and a rectification zone associated with at least one hydroisomerization reaction zone, the hydroisomerization reaction zone being at least partially internal to the distillation zone and comprising at least one catalytic bed, in which hydroisomerization of at least a portion of 1-butene is carried out in the presence of a hydroisomerization catalyst and at least one gas stream comprising hydrogen, such that an effluent rich in isobutene leaves the distillation zone overhead and an effluent rich in 2-butenes leaves the bottom, the process being characterized in that each catalytic bed in the internal portion of the hydroisomerization zone is traversed by an ascending co-current of the gas stream and liquid and is substantially out of contact with the distillation vapor.

This is a division of application Ser. No. 08/774,841 filed Dec. 27,1996 now U.S. Pat. No. 5,888,355.

FIELD OF THE INVENTION

This invention concerns the hydroisomerisation of at least a portion ofthe 1-butene contained in a feed, the major portion of which isconstituted by olefinic hydrocarbons, including isobutene, also 1-buteneand 2-butenes in a ratio which substantially corresponds to thethermodynamic equilibrium.

BACKGROUND OF THE INVENTION

This disclosure also describes apparatus and a hydrogenation processwhich are subjects of related applications. The apparatus of theinvention can be applied to various catalytic reactions, which may beequilibrated or complete, in which at least one of the products of thereaction can be separated in a pure or diluted state by distillationunder temperature and pressure conditions which are close to those ofthe reaction, and more particularly to paraffin isomerisation reactionsinvolving reorganisation of the backbone, olefin isomerisation bydisplacement of the double bond (hydroisomerisation) or byreorganisation of the backbone, hydrogenation of unsaturated compoundsto saturated compounds, dehydrogenation of saturated compounds tounsaturated compounds, and any reaction which requires the presence ofhydrogen.

The hydrogenation catalyst can be disposed in the reaction zone usingthe different techniques proposed for carrying out catalyticdistillation. These techniques have principally been developed foretherification reactions, which involve contact between the reactants ina homogeneous liquid phase and the solid catalyst. They are essentiallyof two types. In the first type of technique, the reaction and thedistillation proceed simultaneously in the same physical space, astaught, for example, in International patent application WO-A-90/02603,U.S. Pat. Nos. 4,471,154, 4,475,005, 4,215,011, 4,307,254, 4,336,407,4,439,350, 5,189,001, 5,266,546, 5,073,236, 5,215,011, 5,275,790,5,338,517, 5,308,592, 5,236,663, 5,338,518 and European patents EP-B1-0008 860, EP-B1-0 448 884, EP-B1-0 396 650 and EP-B 1-0 494 550, alsoEuropean patent application EP-A1-0 559 511. In general, then, thecatalyst is in contact with a descending liquid phase generated by thereflux introduced at the top of the distillation zone, and with anascending vapour phase generated by the reboiling vapour introduced atthe bottom of the zone. In the second type, the catalyst is generallydisposed so that the reaction and distillation proceed independently andconsecutively as taught, for example, in U.S. Pat. No. 4,847,430, U.S.Pat. No. 5,130,102 and U.S. Pat. No. 5,368,691, the vapour from thedistillation step not in practice traversing any catalytic bed in thereaction zone.

For every chemical reaction requiring the addition of a foreign gaseousreactant to the distillation feed, this reactant must be introduced intothe reaction zone in a different manner depending on the type oftechnique selected to carry out the catalytic distillation. In the firsttype, the gaseous reactant can simply be added to the distillationvapour at any level, but always before penetration into the reactionzone, generally substantially at the inlet to at least one catalytic bedof the reaction zone. In the second type, the gaseous reactant must beintroduced in a manner appropriate to the options selected to impose acirculation direction on the liquid and gas in the catalyst bed.

It thus appears that, for a reaction which is carried out in thepresence of a solid catalyst between a liquid phase reactant and agaseous reactant which is only slightly soluble in the liquid, such asthe hydrogenation of unsaturated hydrocarbons mixed with otherhydrocarbons, the option which consists of carrying out catalyticdistillation by avoiding passing the distillation vapour over thecatalyst and causing the liquid and the gaseous reactant to circulate inan ascending co-current in the catalytic bed is the most efficient. Itappears that the pressure drop across the catalytic bed(s) in the firsttype does not produce an intimate mixture of the liquid and gas streamcontaining hydrogen. In effect, in that type of technique, where thereaction and the distillation proceed simultaneously in the samephysical space, the liquid descends through the catalytic bed inrivulets, i.e., in streams of liquid. The gaseous fraction containingthe vaporised fraction of the feed and the gas stream containinghydrogen ascend through the catalytic bed in columns of gas. In thisdisposition, the entropy of the system is greater and the pressure dropacross the catalytic bed(s) is lower. Thus operating in accordance withthe first type does not encourage dissolution of hydrogen in the liquidphase comprising the unsaturated compound(s).

The second type, comprising a specific apparatus for distributing theliquid fraction to be hydrogenated and the gas stream comprisinghydrogen, where said liquid fraction and said gas stream traverse thecatalytic bed in an ascending co-current, can effect the hydrogenationreaction essentially in the absence of the gaseous fraction of the feedand under conditions under which the pressure drop across the catalyticbed(s) is the greatest. Increasing the pressure drop because of thespecific apparatus of the invention can increase the solubility of thehydrogen in the liquid phase and can thus encourage hydrogenation in theliquid fraction.

In addition, because of the recognised toxicity of benzene and olefins,which are unsaturated compounds, the general tendency is to reduce theamount of these constituents in gasoline. Benzene has cancer-causingproperties and thus any possibility of it polluting the ambient air hasto be limited as much as possible, particularly by practically excludingit from automobile fuels. In the United States, reformulated gasolinecannot contain more than 1% of benzene; in Europe, while therestrictions are not yet as severe, it has been recommended that thisvalue be slowly approached. Further, olefins have been recognised asbeing among the most reactive hydrocarbons in the cycle of photochemicalreactions with oxides of nitrogen, which occur in the atmosphere andwhich lead to the formation of ozone. An increase in the ozoneconcentration in air may be a source of respiratory problems. Areduction in the amount of olefins contained in gasoline, in particularthe lightest olefins which have the greatest tendency to volatiliseduring manipulation of the gasoline, is thus desirable.

The benzene content of a gasoline is largely dependent on that of thereformat component of that gasoline. The reformat results from catalytictreatment of naphtha for the production of aromatic hydrocarbons whichprincipally contains 6 to 9 carbon atoms per molecule and in which thevery high octane number provides the gasoline with antiknock properties.Because of the toxicity described above, it is thus necessary to reducethe benzene content of the reformat by as much as possible. A number ofmethods can be envisaged.

The first method consists of limiting the amount of benzene precursorssuch as cyclohexane and methylcyclopentane in the naphtha constitutingthe feed to a catalytic reforming unit. This solution substantiallyreduces the benzene content in the effluent from the reforming unit butis not sufficient in itself when the content must be reduced to as lowas 1%. A second method consists of eliminating a light fraction of thereformat containing benzene by distillation. This solution produces aloss of the order of 15% to 20% of hydrocarbons which could be used inthe gasoline. A third method consists of extracting the benzene presentin the effluent from a reforming unit. A number of techniques can, inprinciple, be applied: solvent extraction, extractive distillation,adsorption, None of these techniques can be applied on an industrialscale, as none of them can selectively extract benzene in an economicalmanner. A fourth method consists of chemically transforming the benzeneto convert it to a constituent which does not suffer legal limits.Alkylation by ethylene, for example, transforms the benzene to mainlyethylbenzene. This operation, however, is expensive because of theoccurrence of secondary reactions which necessitate separation stepswhich use a great deal of energy.

Benzene of a reformat can also be hydrogenated to cyclohexane. As it isimpossible to selectively hydrogenate benzene in a mixture ofhydrocarbons which also contains toluene and xylenes, it is thusnecessary first to fractionate the mixture to isolate a cut whichcontains only benzene, which can then be hydrogenated. A process hasalso been described in which the benzene hydrogenation catalyst isincluded in the rectification zone of a distillation column whichseparates benzene from other aromatics (Benzene Reduction--Kerry Rockand Gary Gildert, CDTECH--1994--Conference on Clean Air ActImplementation and Reformulated Gasoline--October 1994), which can saveon apparatus.

Further, isobutene for polymerisation must be more than 99% pure and canonly contain traces of 1-butene and 2-butenes (several tens of parts permillion by weight, ppm). If the amount of impurities in the isobutene istoo high, the polymers obtained are of poorer quality and thepolymerisation yield is lower. Thus other olefinic hydrocarbonscontaining 4 carbon atoms per molecule must be eliminated from ahydrocarbon cut containing isobutene. 1-butene and isobutene have veryclose boiling points. Separation by distillation is only possible usingdrastic measures. The other olefinic hydrocarbons containing 4 carbonatoms can be separated from the isobutene by distillation. The principalproblem in the production of high purity isobutene is thus separation of1-butene from isobutene. A number of methods can be used to carry outsuch a separation.

The first method consists of extraction using sulphuric acid: isobuteneis selectively hydrated and then regenerated by treating the aqueousphase. If the temperature and concentration are controlled properly,such a process can produce isobutene with high purity. However, theyield normally does not exceed 90%, extraction is not complete anddimers and oligomers are formed which lead to the formation of toxicacid sludge. The second method consists of cracking the methyl ether oftertio-butyl alcohol (MTBE): isobutene is extracted from the C₄ cut byreacting it with methanol to form MTBE. The MTBE is then cracked tomethanol and isobutene on an acid catalyst. The yield can be at least96%. The isobutene produced is of high purity but the dimethyletherwhich can be formed during cracking must be removed. The third possiblemethod is dehydration of tertiary butyl alcohol (TBA). In the precedingoperation, the methanol can be replaced by water, leading to theproduction of TBA. Isobutene is then recovered by dehydration of theTBA. This method is practically never used, primarily because TBA isclosely linked to the propylene oxide market. Using those processes, TBAcan be a by-product of propylene oxide.

U.S. Pat. No. 2,403,672 describes a process for the separation ofisobutene from a mixture of isobutene and 1-butene which comprisesintroducing the mixture into an isomerisation and fractionation zone inwhich the isomerisation catalyst also acts as a packing which carriesout the distillation function. This solution has the major disadvantageof not having good distillation efficiency and thus has a mediocrecapacity for separating the isobutene from the 1-butene. In thistechnique, the reaction and distillation proceed simultaneously in thesame physical space. The catalyst is in contact with a descending liquidphase generated by the reflux introduced at the top of the distillationzone, and with an ascending vapour phase generated by the reboil vapourintroduced to the bottom of the zone.

The invention concerns a reactive distillation apparatus comprising adistillation zone which comprises a stripping zone and a rectificationzone associated with a reaction zone, at least a portion of which isinternal to said distillation zone, and comprising at least onecatalytic bed in which a feed is transformed in the presence of acatalyst and at least one gas stream comprising hydrogen, said apparatusbeing characterized in that each catalytic bed in the internal portionof said reaction zone is traversed by an ascending co-current of saidgas stream and liquid.

The apparatus of the invention generally comprises:

at least one means for distributing the major portion of the liquid fromthe bottom to the top through the catalyst;

at least one means for circulating the major portion of the distillationvapour from the bottom to the top through the catalytic bed such thatsaid vapour is not in practice in contact with the catalyst; and

at least one means for distributing the major portion of the gas streamfrom the bottom towards the top through the catalyst.

The feed supplied to the distillation zone is generally introduced intosaid zone to at least one level in said zone, preferably principally toa single level of said zone.

The distillation zone generally comprises at least one column providedwith at least one distillation contact means selected from the groupformed by plates, loose packing and structured packing, as is known tothe skilled person, and is such that the total overall efficiency isgenerally at least 5 theoretical levels. In known cases in whichoperation of a single column causes problems, it is generally preferableto divide the zone so as to use at least two columns which form saidzone when placed end to end, i.e., the rectification zone, possibly thereaction zone and the stripping zone are distributed over the columns.In practice, when the reaction zone is at least partially internal tothe distillation zone, the rectification or stripping zone, preferablythe stripping zone, is generally in at least one column which isdifferent to the column comprising the internal portion of the reactionzone.

The means for circulating the distributing vapour from the bottomtowards the top through the catalytic bed passes the reaction zone levelwhere the catalytic bed is located, i.e., it is generally located in thecatalytic bed, but it can also be located at the edge of said catalyticbed.

The reaction zone generally comprises at least one catalytic bed,preferably 2 to 6, and more preferably 2 to 4 catalytic bed(s); when atleast two catalytic beds are incorporated into the reaction zone, thesetwo beds may be separated by at least one distillation contact means.

The apparatus of the invention is generally such that the flow of liquidto be transformed is in a co-current with the gas stream comprisinghydrogen and such that the distillation vapour does not in practice passthrough any catalytic bed of the internal portion of the reaction zone(meaning that, in practice, said vapour is separated from said liquid),for any catalytic bed in the internal portion of the reaction zone. Inall cases of this second type of technique, each catalytic bed of theportion of the reaction zone which is in the distillation zone isgenerally such that the gas stream comprising hydrogen and the liquidstream which will react circulates through said bed in a co-current,which is generally ascending, even if overall in the catalyticdistillation zone the gas stream comprising hydrogen and the liquidstream which will react circulates in counter-current mode. Such systemsgenerally comprise at least one apparatus for distributing liquid whichcan, for example, be a liquid distributor, in each catalytic bed in theinternal portion of the reaction zone. Nevertheless, provided that thetechnologies used in the process of the invention have been designed forcatalytic reactions between liquid reactants, without modification theyare not suitable for a catalytic reaction in which one of the reactants,hydrogen, is in the gaseous state. For each catalytic bed of theinternal portion of the reaction zone, it is thus generally necessary toadd an apparatus for introducing a gas stream comprising hydrogen, usingthe techniques described below, for example.

Thus for each catalytic bed in the internal portion of the reactionzone, the internal portion of the reaction zone comprises at least onemeans for distributing liquid, generally located below said catalyticbed, and at least one means for introducing a gas stream, generallylocated below or in said catalytic bed, preferably in the latter caseclose to the liquid introduction means. In one technique, the means forintroducing a gas stream into each catalytic bed is identical to themeans for distributing liquid in the catalytic bed, i.e., there is ameans for introducing gas into the liquid upstream of the means fordistributing liquid (with respect to the direction of circulation of theliquid). In practice and in current parlance, this means that gas isinjected into the liquid upstream of the liquid distribution means. Inanother technique, the means for introducing a gas stream is locatedsubstantially at the level of the liquid distributing means, the gas andliquid being introduced separately into the catalytic bed. In a furthertechnique, the means for introducing a gas stream is located below or inthe catalytic bed, preferably not far from the liquid distributionmeans.

Further, in one embodiment of the invention, the apparatus of theinvention is such that the major portion of said gas stream is hydrogen,the major portion of the hydrogen, and preferably almost all thereof,originating from external the distillation zone.

The apparatus of the invention is generally such that, for the portionof the reaction zone which is internal to the distillation zone, thefeed from the reaction zone is drawn off at a draw-off level andrepresents at least a portion, preferably the major portion, of theliquid flowing in the distillation zone, preferably flowing in therectification zone and more preferably flowing to an intermediate levelof the rectification zone, the effluent from the reaction zone being atleast in part, preferably a major part, re-introduced into thedistillation zone substantially in the proximity, i.e., generallysubstantially at the same height or just above or just below, usually atthe same height or just below, i.e., located at a distance correspondingto a height which is in the range 0 to 4 theoretical plates from adraw-off, preferably from said draw-off, to ensure continuity ofdistillation. Thus for the portion of the reaction zone which isinternal to the distillation zone, the liquid is drawn off naturally byflow in the portion of the reaction zone which is internal to thedistillation zone and re-introduction of the liquid to the distillationzone also occurs naturally by flow of liquid from the portion of thereaction zone which is internal to the distillation zone.

In general, the apparatus of the invention comprises 1 to 4 draw-off(s)which supply the external portion of the reaction zone, when thereaction zone is not completely internal to the distillation zone. Ingeneral, the liquid which will react, either partially or completely,circulates first in the external portion of the reaction zone then inthe internal portion of said zone. Two cases are then possible. In thefirst case, the external portion of the reaction zone is supplied by asingle draw-off and then, if said portion comprises more than tworeactors, these are disposed in series or in parallel. In the secondcase, which is preferred, the external portion of the reaction zone issupplied by at least two draw-offs.

In one of the preferred embodiments of the invention, the apparatus ofthe invention is such that the reaction zone is completely internal tothe distillation zone.

In a preferred embodiment of the apparatus of the invention, thecatalyst is disposed in the reaction zone as described in the basicapparatus defined in U.S. Pat. No. 5,368,691, and arranged such thateach catalytic bed is supplied by the gas stream containing hydrogen,regularly distributed at its base, using one of the techniques describedbelow, for example. Using this technique, if the distillation zonecomprises a single column and if the reaction zone is completely insidesaid column, the catalyst comprised in each catalytic bed, which isinternal to the distillation zone, is then in contact with an ascendingliquid phase generated by the reflux introduced to the top of thedistillation zone, and with hydrogen which circulates in the samedirection as the liquid; contact with the vapour phase from distillationis avoided by passing at least one specially provided chimney throughthe distribution zone.

The invention also concerns a process for the treatment of a feed, amajor portion of which is constituted by hydrocarbons containing atleast 5 and preferably 5 to 9 carbon atoms per molecule, and comprisingat least one unsaturated compound containing at most six carbon atomsper molecule including benzene, and treating said feed in a distillationzone comprising a stripping zone and a rectification zone, associatedwith a hydrogenation reaction zone which is at least partially internalto said distillation zone, in which at least a portion, preferably themajor portion, of the unsaturated compounds containing at most sixcarbon atoms per molecule, i.e., containing up to and including sixcarbon atoms per molecule and contained in the feed are hydrogenated inthe presence of a hydrogenation catalyst and at least one gas streamcomprising hydrogen, preferably as the major portion, to cause aneffluent which is depleted in unsaturated compounds containing at mostsix carbon atoms per molecule to leave overhead of the distillation zoneand an effluent which is depleted in unsaturated compounds containing atmost six carbon atoms per molecule to leave from the bottom of thedistillation zone, characterized in that each catalytic bed of theinternal portion of the hydrogenation zone is traversed by an ascendingco-current of said gas stream and liquid and the catalytic bed is not inpractice traversed by the distillation vapour.

The hydrogenation reaction zone at least partially hydrogenates thebenzene present in the feed, generally in such a way that theconcentration of benzene in the overhead effluent is at most equal to aset concentration, and said reaction zone hydrogenates at least aportion, preferably the major portion, of each unsaturated compoundcontaining at most six carbon atoms per molecule (other than benzene)which maybe present in the feed.

The process of the invention preferably includes the use of theapparatus of the invention.

The distillation zone and the characteristics of the gas stream, thereaction zone, etc. . . . , have been described above with respect tothe apparatus of the invention.

In one implementation of the process of the invention, the effluent fromthe bottom of the distillation zone is mixed with the overhead effluentfrom said zone, In this case, after any stabilisation which may benecessary, the mixture obtained is used as a fuel either directly or byincorporation of fuel fractions.

To carry out hydrogenation using the process of the invention, thetheoretical molar ratio of hydrogen which is necessary for the desiredconversion of benzene is 3. The quantity of hydrogen injected before orinto the hydrogenation zone is optionally in excess with respect to thisstoichiometry, more so when in addition to the benzene present in thefeed, each unsaturated compound containing at most six carbon atoms permolecule present in the feed must be hydrogenated. If the conditions aresuch that there is an excess of hydrogen, the excess hydrogen canadvantageously be recovered using one of the techniques described below,for example. As an example, the excess hydrogen which leaves thedistillation zone overhead is recovered then injected upstream of thecompression steps associated with a catalytic reforming unit, mixed withthe hydrogen from said unit, said unit preferably operating at lowpressure (i.e., generally at a pressure of less than 8 bars). Thisexcess hydrogen can also be recovered then compressed and used again inthe reaction zone.

The major portion, preferably almost all, of the hydrogen used in thereaction zone of the invention generally originates from external thedistillation zone. It can originate from any source which produceshydrogen of at least 50% purity by volume, preferably at least 80%purity by volume and more preferably at least 90% purity by volume. Asan example, hydrogen originating from catalytic reforming processes,from PSA (pressure swing adsorption), electrochemical generation, steamcracking or steam reforming can be used.

The operating conditions in the hydrogenation zone in the process of theinvention are linked to the operating conditions used for distillation.Distillation is carried out at a pressure which is generally in therange 2 to 20 bars, preferably in the range 4 to 15 bars, and morepreferably in the range 4 to 10 bars (1 bar=10⁵ Pa), with a reflux ratioin the range 1 to 10, preferably in the range 3 to 6. The temperature atthe head of the zone is generally in the range 40° C. to 180° C. and thetemperature at the bottom of the zone is generally in the range 120° C.to 280° C. The hydrogenation reaction is carried out under conditionswhich are most generally intermediate between those established overheadand at the bottom of the distillation zone, at a temperature which is inthe range 100° C. to 200° C., preferably in the range 120° C. to 180°C., and at a pressure which is in the range 2 to 20 bars, preferably inthe range 4 to 10 bars. The liquid which is hydrogenated is suppliedwith hydrogen, the flow rate of which depends on the concentration ofbenzene in said liquid and, more generally, of the unsaturated compoundscontaining at most six carbon atoms per molecule in the feed from thedistillation zone. It is generally at least equal to the flow rate whichcorresponds to the stoichiometry of the hydrogenation reactions takingplace (hydrogenation of benzene and other unsaturated compoundscontaining at most six carbon atoms per molecule comprised in thehydrogenation feed) and at most equal to the flow rate which correspondsto 10 times the stoichiometry, preferably less than six times thestoichiometry, and more preferably less than 3 times the stoichiometry.

When the hydrogenation zone includes a portion which is external to thedistillation zone, the catalyst disposed in said external portion canuse any technique which is known to the skilled person, under operatingconditions (temperature, pressure . . . ) which are or are notindependent, preferably independent, of the operating conditions in thedistillation zone. In the portion of the hydrogenation zone which isexternal to the distillation zone, the operating conditions aregenerally as follows. The pressure required for this hydrogenation stepis generally in the range 1 to 60 bars absolute, preferably in the range2 to 50 bars and more preferably in the range 5 to 35 bars. Theoperating temperature in the hydrogenation zone is generally in therange 100° C. to 400° C., preferably in the range 120° C. to 350° C.,more preferably in the range 140° C. to 320° C. The space velocity insaid hydrogenation zone, calculated with respect to the catalyst, isgenerally in the range 1 to 50 h⁻¹, more particularly in the range 1 to30 h⁻¹ (volume of feed per volume of catalyst per hour). The hydrogenflow rate corresponding to the stoichiometry of the hydrogenationreactions taking place is in the range 0.5 to 10 times saidstoichiometry, preferably in the range 1 to 6 times said stoichiometryand more preferably in the range 1 to 3 times said stoichiometry.However, the temperature and pressure conditions can also be withinthose established at the head and bottom of the distillation zone,without departing from the scope of the invention.

More generally, whatever the position of the hydrogenation zone withrespect to the distillation zone, the catalyst used in the hydrogenationzone of the invention generally comprises at least one metal selectedfrom the group formed by nickel and platinum, used as they are orpreferably deposited on a support. The metal is generally in its reducedform for at least 50% by weight of its full quantity. However, any otherhydrogenation catalyst which is known to the skilled person can also beused.

When platinum is used, the catalyst can advantageously contain at leastone halogen in a proportion which is in the range 0.2% to 2% by weightwith respect to the catalyst. Chlorine or fluorine is preferably used,or a combination of the two in a proportion which is in the range 0.2%to 1.5% with respect to the total catalyst weight. When a catalystcontaining platinum is used, a catalyst is generally used in which theaverage size of the platinum crystallites is below 60×10⁻¹⁰ m,preferably less than 20×10⁻¹⁰ m and more preferably less than 10×10⁻¹⁰m. Further, the overall proportion of the platinum is generally in therange 0.1% to 1% with respect to the total catalyst weight, preferablyin the range 0.1% to 0.6%.

When nickel is used, the proportion of nickel is in the range 5% to 70%with respect to the total weight of the catalyst, more particularly inthe range 10% to 70% and preferably in the range 15% to 65%. Further, acatalyst is generally used in which the average size of the nickelcrystallites is less than 100×10⁻¹⁰ m, preferably less than 80×10⁻¹⁰ m,and more preferably less than 60×10⁻¹⁰ m.

The support is generally selected from the group formed by alumina,silica-aluminas, silica, zeolites, activated charcoal, clays, aluminouscements, rare earth oxides and alkaline-earth oxides, used alone or as amixture. Preferably, a support based on alumina or silica is used, witha specific surface area which is in the range 30 to 300 m² /g,preferably in the range 90 to 260 m² /g.

This particularly claimed invention concerns a process for the treatmentof a feed comprising, as its major portion, olefinic hydrocarbonscontaining 4 carbon atoms per molecule, including isobutene, also1-butene and 2-butenes in a ratio which substantially corresponds to thethermodynamic equilibrium, in which said feed is treated in adistillation zone comprising a stripping zone and a rectification zoneassociated with a hydroisomerisation reaction zone, said reaction zonebeing at least partially internal to said distillation zone andcomprising at least one catalytic bed, in which hydroisomerisation of atleast a portion and preferably the major portion of 1-butene is carriedout in the presence of a hydroisomerisation catalyst and a gas streamcomprising hydrogen, preferably as its major portion, such that aneffluent which is rich in isobutene, generally of high purity, leavesthe distillation zone overhead and an effluent which is depleted inisobutene leaves the bottom, said process being characterized in thateach catalytic bed in the internal portion of the hydroisomerisationzone is traversed by an ascending co-current of said gas stream andliquid and is not in practice traversed by distillation vapour. Theprocess can be used to produce high purity isobutene.

The process of the invention preferably includes the use of theapparatus of the invention.

The feed supplied to the distillation zone is generally introduced intosaid zone to at least one level of said zone, preferably principally toa single level of said zone. It is in a ratio which substantiallycorresponds to the 1-butene/2-butenes thermodynamic equilibrium onintroduction. In a preferred implementation of the process of theinvention, the feed is obtained from a cut with a major portioncomprised of olefinic hydrocarbons containing 4 carbon atoms permolecule, including isobutene and 1-butene, by treatment of said cut ina first hydroisomerisation zone, which is generally independent of theoptional portion of the hydroisomerisation reaction zone which isexternal to the associated distillation zone, the major portion of theeffluent from said first hydroisomerisation zone acting as the feed,which is principal or secondary according to the definitions given belowin the text, which supplies the distillation zone. If the feed includespolyunsaturated compounds, usually dienes and/or acetylenes, thecompounds are preferably transformed to butenes in the firsthydroisomerisation zone before introduction into the distillation zone.However, any other technique which can produce a feed in which the1-butene and 2-butenes are in a ratio which substantially corresponds tothe thermodynamic equilibrium from a cut with olefin C₄ hydrocarbons asits major portion is also within the scope of the invention.

The first optional hydroisomerisation reaction zone located upstream ofthe distillation-reaction zone effects at least partial selectivehydrogenation of the polyunsaturated compounds, usually dienes such asbutadiene, in addition to hydroisomerisation of at least a portion ofthe 1-butene to 2-butenes. It generally comprises at least one catalytichydroisomerisation bed comprising a hydroisomerisation catalyst,preferably with 1 to 4 catalytic beds; when at least two catalytic bedsare incorporated into the reaction zone, these two beds are preferablydistributed over at least two reactors, in series or in parallel,preferably in series. As an example, said first reaction zone comprisesa single reactor containing at least one catalytic bed, preferably onlyone. In a preferred implementation of the process of the presentinvention, said first reaction zone comprises two reactors which aregenerally in series each comprising at least one catalytic bed,preferably only one. When the reaction zone comprises at least tworeactors, any recycling of at least a portion of the effluent from atleast one of the reactors in the first reaction zone to the first zoneis generally made to the inlet of one reactor, preferably to saidreactor, preferably before injection of the gaseous compound comprisinghydrogen. It is also possible to recycle around the first zone itself,i.e., generally to the inlet of the first reactor to said zone,preferably before injection of the gaseous compound comprising hydrogen;as an example, with two reactors, at least a portion of the effluentfrom the second reactor is recycled to the inlet to the first reactor.This can advantageously reduce the concentration of polyunsaturatedcompounds in the effluent from the first reaction zone.

The operating conditions of the first hydroisomerisation zone, whenpresent, are generally as follows: the catalyst is identical to thecatalyst from the hydroisomerisation zone which will be described below.The pressure is generally in the range 4 to 40 bar (1 bar=0.1 MPa),preferably in the range 6 to 30 bar. The temperature is generally in therange 10° C. to 150° C., preferably in the range 20° C. to 100° C. TheH₂ /hydrocarbons molar ratio is generally adjusted so as to obtainpractically complete conversion of the polyunsaturated compounds such asbutadiene and sufficient isomerisation of 1-butene to 2-butenes withlimited alkane formation.

The hydroisomerisation reaction zone associated with the distillationzone generally comprises at least one catalytic hydroisomerisation bedcomprising a hydroisomerisation catalyst, preferably 2 to 4, and morepreferably 2 to 6 catalytic beds; when at least two catalytic beds areincorporated into said distillation zone, these two beds are preferablyseparated by at least one distillation contact means. Thehydroisomerisation reaction zone at least partially hydroisomerises atleast a portion, preferably the major portion, of the 1-butene presentin the feed to 2-butenes (cis and trans), generally such that the1-butene concentration in the overhead effluent from the distillationzone is a maximum of a certain value.

The distillation zone used in the process of the invention is identicalto that described above.

In a preferred implementation of the process of the invention, inaddition to supplying the distillation zone with the principal feed, itis supplied with a secondary feed (secondary with respect to theprincipal feed) which may or may not originate from a hydroisomerisationreaction zone such as the first, optional, hydroisomerisation reactionzone, and may or may not be independent of the supply of principal feedto the distillation zone. The secondary feed is generally a C₄ cutcontaining at least isobutene, also 1-butene and 2-butenes in a ratiowhich substantially corresponds to the thermodynamic equilibrium, andgenerally originates from a steam cracking process such as a crude C₄cut or the 1-raffinate, or from catalytic cracking; generally andpreferably, the secondary feed is a C₄ cut which is essentially free ofpolyunsaturated compounds and the 1-butene content is lower than the1-butene content in the principal feed. If the amount of unsaturatedcompounds in the secondary feed is high, the feed is preferably treatedin a selective hydrogenation zone before its entry into the distillationzone.

When the principal feed is introduced at a single introduction level,the secondary feed is generally introduced into the distillation zone toat least one introduction level, preferably to a single introductionlevel, said introduction level depending on the composition of thesecondary feed. Thus in a first example, the secondary feed can be veryrich in isobutene and contain less than 1.5 times the 1-butene containedin the principal feed, in which case the secondary feed is preferablyintroduced at a single level generally located above the level at whichthe principal feed is introduced. In a second example, the secondaryfeed is practically free of 1-butene, in which case the secondary feedis preferably introduced to a single level generally located below thelevel at which the principal feed is introduced. It is also possible tomix the principal feed before its entry into the distillation zone withthe secondary feed.

The hydroisomerisation reaction zone associated with the distillationzone generally comprises at least one catalytic hydroisomerisation bed,preferably 2 to 6 and more preferably 2 to 4 catalytic beds; when atleast two catalytic beds are incorporated into the distillation zone,these two beds are optionally separated by at least one distillationcontact means. The hydroisomerisation reaction zone at least partiallyhydroisomerises at least a portion, preferably the major portion, of the1-butene present in the feed to 2-butenes (cis and trans), generallysuch that the 1-butene concentration in the overhead effluent from thedistillation zone is a maximum of a pre-set value.

The process of the invention is generally such that the flow of theliquid to be hydroisomerised is in a co-current with the flow of the gasstream comprising hydrogen in each catalytic bed of the internal portionof the hydroisomerisation zone, and such that the distillation vapourdoes not in practice traverse any catalytic bed in the internal portionof the reaction zone (meaning that in practice, the vapour is separatedfrom the liquid to be hydroisomerised). Each catalytic bed in theportion of the reaction zone which is internal to the distillation zoneis generally such that the gas stream comprising hydrogen and the liquidstream which is to be reacted circulate in a co-current, generallyascending, across the bed, even if overall in the catalytic distillationzone, the gas stream comprising hydrogen and the liquid stream to bereacted circulate in a counter-current. Such systems generally compriseat least one liquid distribution apparatus which can, for example, be aliquid distributor, for each catalytic bed in the internal portion ofthe reaction zone. The distribution apparatus for the gas streams andfor liquid distribution have been described above.

The process of the invention is generally such that in all parts of thehydroisomerisation reaction zone, whether internal or, optionally,external, the feed is drawn off at the height of a draw-off andrepresents at least a portion, preferably the major portion, of theliquid (reflux) flowing in the distillation zone, preferably flowing inthe rectification zone and more preferably flowing in an intermediatelevel of the rectification zone, the effluent from thehydroisomerisation reaction zone being at least in part, preferably themajor part, re-introduced into the distillation zone, so as to ensurecontinuity of distillation. For the optional portion of the reactionzone which is external of the distillation zone, re-introduction of theeffluent from the distillation zone is effected substantially in theproximity, i.e., generally substantially at the same height or justabove or just below, generally at the same height or just above, i.e.,located at a distance corresponding to a height which is in the range 0to 4 theoretical plates from a draw-off, preferably from said draw-off,to ensure continuity of distillation. For the portion of the reactionzone which is internal to the distillation zone, liquid (reflux)draw-off occurs naturally by flow in the portion of the reaction zonewhich is internal to the distillation zone, and re-introduction of theeffluent to the distillation zone also occurs naturally by flow ofliquid from the internal reaction zone to the distillation zone.

In general, when the hydroisomerisation zone is not completely internalto the distillation zone, the process of the invention comprises 1 to 6,preferably 2 to 4 draw-offs, which supply the external portion of thehydroisomerisation zone. In such a case, the liquid to behydroisomerised, either partially or completely, circulates first in theexternal portion of the hydroisomerisation zone then in the internalportion of that zone. Two cases are then possible. In the first case,the external portion of the reaction zone is supplied by a singledraw-off and thus, if said portion comprises more than two reactors,these are disposed in series or in parallel. In the second case, whichis preferred, the external portion of the hydroisomerisation zone issupplied by at least two draw-offs. A portion of the external portion ofthe hydroisomerisation zone which is supplied by a given draw-off, ifthe external portion comprises at least two draw-offs, generallycomprises at least one reactor, preferably a single reactor. If saidportion of the external portion comprises at least two reactors, eachreactor which is external to the distillation zone is generally suppliedby a single draw-off, preferably associated with a singlere-introduction level, said draw-off being distinct from the draw-offwhich supplies the other reactor(s).

In preferred implementation of the invention, the process of theinvention is such that the hydrogenation zone is completely internal tothe distillation zone.

The major portion of the hydrogen used for hydroisomerisation of1-butene, preferably almost all thereof, originates from external thedistillation zone. It can originate from any source which produceshydrogen in at least 50% purity by volume, preferably at least 80%purity by volume and more preferably at least 90% purity by volume. Asan example, hydrogen originating from catalytic reforming processes,from PSA (pressure swing adsorption), electrochemical generation, steamcracking or steam reforming can be used.

The operating conditions in the portion of the hydroisomerisation zoneinternal to the distillation zone are linked to the operating conditionsused for distillation. Distillation is generally carried out in a mannerwhich minimises the quantity of isobutene in the bottom product tomaximise the yield of isobutene from the process and minimise thequantity of 2-butenes and 1-butene in the overhead product to producehigh purity isobutene overhead. It is carried out at a pressure which isgenerally in the range 2 to 30 bars, preferably in the range 4 to 15bars, and more preferably in the range 4 to 10 bars, with a reflux ratioin the range 1 to 30, preferably in the range 5 to 20. The temperatureat the head of the zone is generally in the range 0° C. to 200° C. andthe temperature at the bottom of the zone is generally in the range 5°C. to 250° C. The hydroisomerisation reaction is carried out underconditions which are most generally intermediate between thoseestablished overhead and at the bottom of the distillation zone, at atemperature which is in the range 20° C. to 150° C., preferably in therange 40° C. to 80° C., and at a pressure which is in the range 2 to 30bars, preferably in the range 4 to 15 bars, and more preferably in therange 4 to 10 bars. The liquid which is hydroisomerised is supplied witha gas stream comprising hydrogen, preferably as the major portion.

When the hydroisomerisation zone includes a portion which is external tothe distillation zone, the catalyst disposed in said external portioncan use any technique which is known to the skilled person, underoperating conditions (temperature, pressure . . . ) which are generallyindependent of the operating conditions in the distillation zone. In theoptional portion of the hydroisomerisation zone which is external to thedistillation zone, the operating conditions are generally as follows.The pressure required for this hydroisomerisation step is generally inthe range of about 1 to 40 bars absolute, preferably in the range ofabout 2 to 30 bars and more preferably in the range of about 4 to 25bars. The operating temperature in the hydroisomerisation zone isgenerally in the range of about 20° C. to 150° C., preferably in therange of about 40° C. to 100° C., more preferably in the range of about40° C. to 80° C. The space velocity in said hydroisomerisation zone,calculated with respect to the catalyst, is generally in the range ofabout 1 to 100 h⁻¹, more particularly in the range of about 4 to 50 h⁻¹(volume of feed per volume of catalyst per hour). The correspondinghydrogen flow rate is such that the H₂ /hydrocarbons molar ratio onentering the hydroisomerisation zone is preferably at least about 10⁻⁵.This ratio is usually about 10⁻⁵ to about 3 and often about 10⁻³ toabout 1. However, the temperature and pressure conditions can also be inthe range which is established at the head and bottom of thedistillation zone, without departing from the scope of the invention.

In order to carry out hydroisomerisation using the process of theinvention, the theoretical molar ratio of hydrogen necessary for thedesired conversion of 1-butene in the reaction zone associated with thedistillation zone is such that the H₂ /hydrocarbons molar ratio onentering said zone is at least 10⁻⁵. This molar ratio can be optimisedso that all the hydrogen is consumed in the hydroisomerisation reactionto avoid the need for a hydrogen recovery apparatus at the outlet to thereaction zone, and so that the secondary hydrogenation reactions can beminimised to maximise the isobutene yield from the process and finally,such that there is sufficient hydrogen along the whole length of thereaction zone so that the hydroisomerisation reaction of 1-butene to2-butenes can take place. However, if these conditions are such thatthere is an excess of hydrogen, the excess hydrogen can advantageouslybe recovered using one of the techniques described above, for example.As an example, the excess hydrogen leaving the distillation zoneoverhead is recovered, then injected upstream of the compression stagesassociated with a catalytic reforming unit, mixed with the hydrogen fromsaid unit, said unit preferably operating at low pressure (i.e.,generally a pressure of less than 8 bar). This excess hydrogen can alsobe recovered then compressed and used again in the reaction zone.

When a portion of the hydroisomerisation zone associated with thedistillation zone is external to the distillation zone, the process ofthe invention can isomerise a large portion of the 1-butene to 2-butenesexternal the distillation zone, optionally under different temperatureand/or pressure conditions to those used in the column. The inlettemperature (and, respectively, the outlet temperature) at the draw-offwhich supplies a catalytic bed of the portion of the hydroisomerisationzone which is external to the column is preferably substantiallysimilar, i.e., the difference is substantially less than 10° C. withrespect to the temperature at the height of the draw-off (with respectto the re-introduction level). Similarly, the hydroisomerisationreaction can advantageously be carried out in the portion of thereaction zone which is external the column at a pressure which is higherthan that used internal to the distillation zone. This pressure increasecan thus increase dissolution of the gas stream containing hydrogen inthe liquid phase containing 1-butene to be isomerised.

In such a case, the process of the invention comprises the use of atechnique known as "pumparound" which consists of passing a portion,preferably the major portion, of the liquid (reflux) outside thedistillation zone in an amount which is preferably a factor of more than1, i.e., the flow rate of a catalytic bed in the external portion of thehydroisomerisation zone associated with the distillation zone, said bedbeing supplied at a draw-off with a portion of the liquid effluent(reflux) flowing on the distillation plate associated with said draw-off(i.e., from which said portion of liquid effluent is drawn off) and withat least a portion of the liquid corresponding to recycling the effluentfrom said bed just above or just below or substantially at the samelevel as said draw-off is more than once the flow rate of the liquidflowing on said plate, for example 1.5 times.

More generally, the catalyst used in the hydroisomerisation zone of theprocess of the invention generally comprises at least one metal selectedfrom the group formed by noble metals from group VIII of the periodicclassification of the elements and nickel, i.e., selected from the groupformed by ruthenium, rhodium, palladium, osmium, iridium and platinum,preferably palladium, or nickel, used as it is or, preferably, depositedon a support. At least 50% by weight of the metal is generally in itsreduced form. The noble metal content of the catalyst is generally about0.01% to about 2% by weight. When nickel is used, the proportion ofnickel with respect to the total catalyst weight is in the range 5% to70%, preferably in the range 10% to 70%, and in general, a catalyst isused in which the average size of the nickel crystallites is less than10 nm, preferably less than 8 nm, more preferably less than 6 nm.However, any other hydroisomerisation catalyst which is known to theskilled person can also be selected. Before use, the catalyst isnormally treated with a sulphur compound then by hydrogen. The catalystis generally sulphurated in situ or ex situ such that sulphur ischemisorbed onto at least a portion of the metal. The chemisorbedsulphur encourages the isomerisation of 1-butene to 2-butenes over theisobutene hydrogenation reaction and thus maximises the isobutene yieldof the process.

The hydroisomerisation catalyst support is generally selected from thegroup formed by alumina, silica-aluminas, silica, zeolites, activatedcharcoal, clays, aluminous cements, rare earth oxides and alkaline-earthoxides, used alone or as a mixture. A support based on alumina or silicais preferably used, with a specific surface area which is in the range10 to 300 m² /g, preferably in the range 30 to 70 m² /g.

By way of non limiting example, according to the present invention,commercial catalysts can be used, such as those sold by the companyCatalysts and Chemicals under reference C-31 or those sold by theGirdler Corporation under reference G-55 or, preferably, those sold byProcatalyse under references LD-265, LD-265S, LD-267 and LD-267R.

EXAMPLES

Examples 1 and 2 show the operation of a zone for hydrogenation ofunsaturated compounds comprising at most six carbon atoms per molecule,including benzene, in accordance with the invention (Example 2) and theoperation of a hydrogenation zone which was not in accordance with theinvention (Example 1), with a catalyst which was loosely packed ondistillation plates traversed by the liquid which circulated downwardlyand by a vapour which circulated upwardly in the distillation zone.

Example 1 (comparative)

A metal distillation column with a diameter of 50 mm was used, renderedadiabatic with heating envelopes in which the temperatures wereregulated to produce the temperature gradient established in the column.Over a height of 4.5 m, the column comprised, from head to foot arectification zone composed of 11 sieve plates with downcomers, acatalytic hydrogenating distillation zone and a stripping zone composedof 63 perforated plates. The catalytic hydrogenating distillation zonewas constituted by three reactive plates which in this instance weresieve plates with downcomers, with the weirs raised by 3.5 cm and inwhich the volume between the top of the weir and the plate could bepacked with catalyst. A metal screen placed at the top of the overflowacted as a filter to prevent catalyst particles from being evacuatedwith the liquid leaving the plate.

Each of the three cells was packed with 36 g of nickel catalyst sold byPROCATALYSE with the reference LD 746. 260 g/h of a reformat comprisingessentially hydrocarbons containing at least 5 carbon atoms per moleculewas introduced to the 37^(th) plate in the column counting from thebottom. The reformat composition is shown in Table 1. 18 N1/h ofhydrogen was also introduced to the base of each cell. The column wasstarted by establishing a reflux ratio of 5, regulating the bottomtemperature to 176° C. and the pressure at 7 bars.

At steady state, 138 g/h and 113 g/h respectively of residue and aliquid distillate were recovered. The compositions are given in Table 1.A small portion of the distillate, constituted by the lightesthydrocarbons, was evacuated from the column with excess hydrogen and wasnot accounted for. Analysis of the effluents led to the deduction thatthe degree of hydrogenation of the olefins and benzene in the feed wererespectively 100% and 55% while toluene was unaffected.

Example 2: (in accordance with the invention)

The apparatus of Example 1 was used but the catalytic distillation zonewas of a different design. The catalytic hydrogenating distillation zonein this instance was constituted by three catalytic distillationdoublets, each doublet being itself constituted by a catalytic cell overwhich three perforated plates were mounted. Construction details of acatalytic cell and its disposition in the column are shown schematicallyin the figure. Catalytic cell 1 consists of a cylindrical container witha flat base with an external diameter which is 2 mm smaller than theinternal diameter of the column. At its lower portion above the base, ascreen 2 is provided which acts both as a support for the catalyst andas a distributor for the hydrogen. A catalyst retaining screen 3 isprovided at the top; the height thereof can be varied. Catalyst 4 fillsthe entire volume between these two screens. The catalytic cell receivesthe liquid from upper distillation plate 5 via downcomer 6. Afterpassing through the cell in the upward direction, the liquid isevacuated by overflowing via downcomer 7 and flows onto lowerdistillation plate 8. The vapour from lower plate 8 passes into thecentral chimney 9, which is solid with the cell, by penetrating viaorifices 10 (only one shown in the figure) and leaving below upper plate5 via orifices 11 (only one shown in the figure). Hydrogen is introducedto the base of the catalytic cell via conduit 12 then via orifices 13(six in total) which are distributed around the periphery of the cell,in the immediate vicinity of the base. Seals 14 prevent any hydrogenfrom escaping prior to its arrival at the catalytic bed.

Each of the three cells was packed with 36 g of nickel catalyst sold byPROCATALYSE with the reference LD 746. 260 g/h of the same feed as thatused in Example 1 was introduced to the 37^(th) plate in the columncounting from the bottom. The reformat composition is shown in thesecond column of the Table. 6 N1/h of hydrogen was introduced to thebase of each cell. The column was started by establishing a reflux ratioof 5, regulating the bottom temperature to 176° C. and the pressure at 7bars.

At steady state, 143 g/h and 106 g/h respectively of residue and aliquid distillate were recovered. The compositions are given in Table 1.A small portion of the distillate, constituted by the lightesthydrocarbons, was evacuated from the column with excess hydrogen and wasnot accounted for. Analysis of the effluents led to the deduction thatthe degree of hydrogenation of the olefins and benzene in the feed wererespectively 100% and 87% while toluene was unaffected.

                  TABLE 1                                                         ______________________________________                                        compositions of feed and effluents in the catalytic column                                composition, in % by weight                                                       example 1       example 2                                                            liquid         liquid                                    feed residue distillate residue distillate                                  ______________________________________                                        C5 and lighter                                                                             7.65          10.22         7.36                                   of which: olefins  0.11  0    0                                               C6 44.83  9.55 89.78 12.4  92.59                                              of which: olefins  0.13  0    0                                               :benzene  6.07  0.63  5.45  0.07  1.84                                        :cyclohexane 1.1  8.34  0.34 12.16  0.73                                      C7: 42.55 80.72  78.27  0.05                                                  of which: toluene  4.78 9.1   8.87                                            C8 and heavier  4.97  9.73   9.33                                           olefin conversion                                                                         100%            100%                                                benzene conversion  55%  87%                                                  hydrogen conversion  15%  70%                                               ______________________________________                                    

It can be seen that the process of the present invention produces betterconversion of benzene and better hydrogen conversion.

Examples 3 and 4 illustrate the case of a process of the invention fortreatment of a feed the major portion of which is comprised by olefinichydrocarbons containing 4 carbon atoms per molecule, includingisobutene, 1-butene and 2-butenes in a ratio which substantiallycorresponds to the thermodynamic equilibrium.

Example 3

Hydroisomerisation operations were carried out successively anddiscontinuously on a C₄ distillation cut. The feed was hydroisomerised afirst time. The effluent from the first test was distilled: thedistillation head, representing an intermediate extraction, washydroisomerised. The hydroisomerisation effluent, representing whatwould be re-injected into the column, was distilled. The head from thesecond distillation was hydroisomerised and the effluent from this thirdhydroisomerisation step was distilled.

The hydroisomerisation operations were carried out in a pilot unitprovided with an adiabatic reactor. The reactor was filled with 1.5 l ofcatalyst LD-265 from PROCATALYSE. The catalyst was sulphurated andactivated in situ using the procedure recommended by the supplier of thecatalyst.

Distillation operations were carried out in an adiabatic column with aninternal diameter of 163 mm and a height of 10 m. the column wasconstituted by 4 beds which were 1.78 m high above the feed injectionpoint, filled with a packing sold by SUIZER under the trade name M550Yand 2 beds 1 m high below the feed injection point, filled with Pallrings.

First Hydroisomerisation

The average operating conditions during the test were as follows:

Reactor temperature: 80° C.

Reactor pressure: 20 bar

Residence time: 0.25 h

H₂ /feed molar ratio: 3

Table 2 below shows the compositions of the feed and effluent in thehydroisomerisation reactor operating under the conditions describedabove.

                  TABLE 2                                                         ______________________________________                                                    Feed    Effluent                                                    (weight %) (weight %)                                                       ______________________________________                                        <C.sub.4       0.25      0.23                                                   iC.sub.4  2.98  3.10                                                          iC.sub.4.sup.= 44.90 44.42                                                    C.sub.4 .sup.= 1 26.95  4.26                                                  C.sub.4.sup.== 1,3  0.13  0.00                                                nC.sub.4 11.72 14.41                                                          C.sub.4.sup.= 2trans  8.73 21.37                                              Neo C.sub.5  0.24  0.23                                                       Me cyclo C.sub.3  0.06  0.06                                                  C.sub.4.sup.= 2cis  4.03 11.92                                                >C.sub.4  0.01  0.00                                                        ______________________________________                                    

The legend for the table and the following tables is as follows:

<C₄ : compounds with less than 4 (4 excluded) carbon atoms per molecule(or C₃ ⁻)

iC₄ : isobutene

iC₄ ⁼ : isobutene

C₄ ⁼ 1: 1-butene

C₄ ⁼⁼ 1,3: 1,3-butadiene

nC₄ : normal butane

C₄ ⁼ 2trans: trans 2-butene

Neo C₅ : neopentane (or dimethylpropane)

Me cyclo C₃ : methyl cyclopropane

C₄ ³² 2cis: cis 2-butene

>C₄ : compounds containing more than 4 (4 excluded) carbon atoms permolecule (or C₅ ⁺)

First Distillation

Distillation of the effluent from the above test was carried out underthe following operating conditions:

Column pressure: 4 bar

Reflux ratio (R/D): 20

Feed temperature: 33° C.

Reflux temperature: 32° C.

Column head temperature: 57° C.

Column bottom temperature: 63° C.

Table 3 below shows the compositions of the feed and the overheadeffluent from the distillation column operating under the conditionsdescribed above.

                  TABLE 3                                                         ______________________________________                                                    Feed    Head                                                        (weight %) (weight %)                                                       ______________________________________                                        <C.sub.4       0.23      0.44                                                   iC.sub.4  3.10  6.71                                                          iC.sub.4.sup.= 44.42 83.35                                                    C.sub.4 .sup.= 1  4.26  7.39                                                  C.sub.4.sup.== 1,3  0.00  0.00                                                nC.sub.4 14.41  1.62                                                          C.sub.4.sup.= 2trans 21.37  0.44                                              Neo C.sub.5  0.23 --                                                          Me cyclo C.sub.3  0.06 --                                                     C.sub.4.sup.= 2cis 11.92  0.05                                                >C.sub.4 -- --                                                              ______________________________________                                    

Second Hydroisomerisation

The average operating conditions during the test were as follows:

Reactor temperature: 65° C.

Reactor pressure: 20 bar

Residence time: 0.25.h

H₂ /feed molar ratio: 0.6

Table 4 below shows the compositions of the feed and effluent in thehydroisomerisation reactor operating under the conditions describedabove.

                  TABLE 4                                                         ______________________________________                                                    Feed    Effluent                                                    (weight %) (weight %)                                                       ______________________________________                                        <C.sub.4       0.44      0.39                                                   iC.sub.4  6.71  6.91                                                          iC.sub.4.sup.= 83.35 82.94                                                    C.sub.4 .sup.= 1  7.39  0.81                                                  C.sub.4.sup.== 1,3 -- --                                                      nC.sub.4  1.62  2.09                                                          C.sub.4.sup.= 2trans  0.44  4.44                                              Neo C.sub.5 -- --                                                             Me cyclo C.sub.3 -- --                                                        C.sub.4.sup.= 2cis  0.05  2.42                                                >C.sub.4 -- --                                                              ______________________________________                                    

Second Distillation

Distillation of the effluent from the above test was carried out underthe following operating conditions:

Column pressure: 4 bar

Reflux ratio (R/D): 13.5

Feed temperature: 36° C.

Reflux temperature: 41° C.

Column head temperature: 51° C.

Column bottom temperature: 55° C.

Table 5 below shows the compositions of the feed and the overheadeffluent from the distillation column operating under the conditionsdescribed above.

                  TABLE 5                                                         ______________________________________                                                    Feed    Head                                                        (weight %) (weight %)                                                       ______________________________________                                        <C.sub.4       0.39      0.65                                                   iC.sub.4  6.91 13.71                                                          iC.sub.4.sup.= 82.94 84.82                                                    C.sub.4 .sup.= 1  0.81  0.51                                                  C.sub.4.sup.== 1,3 -- --                                                      nC.sub.4  2.09  0.14                                                          C.sub.4.sup.= 2trans  4.44  0.12                                              Neo C.sub.5 -- --                                                             Me cyclo C.sub.3 -- --                                                        C.sub.4.sup.= 2cis  2.42  0.05                                                >C.sub.4 -- --                                                              ______________________________________                                    

Third Hydroisomerisation

The average operating conditions during the test were as follows:

Reactor temperature: 60° C.

Reactor pressure: 20 bar

residence time: 0.25 to 0.1 h

H₂ /feed molar ratio: 1

Table 6 below shows the compositions of the feed and effluent in thehydroisomerisation reactor operating under the conditions describedabove.

                  TABLE 6                                                         ______________________________________                                                    Feed    Effluent                                                    (weight %) (weight %)                                                       ______________________________________                                        <C.sub.4       0.65      0.57                                                   iC.sub.4 13.71 14.55                                                          iC.sub.4.sup.= 84.82 84.07                                                    C.sub.4 .sup.= 1  0.51  0.03                                                  C.sub.4.sup.== 1,3 -- --                                                      nC.sub.4  0.14  0.22                                                          C.sub.4.sup.= 2trans  0.12  0.38                                              Neo C.sub.5 -- --                                                             Me cyclo C.sub.3 -- --                                                        C.sub.4.sup.= 2cis  0.05  0.18                                                >C.sub.4 -- --                                                              ______________________________________                                    

Third Distillation

Distillation of the effluent from the above test was carried out underthe following operating conditions:

Column pressure: 4 bar

Reflux ratio (R/D): 13.5

Feed temperature: 36° C.

Reflux temperature: 41° C.

Column head temperature: 53° C.

Column bottom temperature: 55° C.

Table 7 below shows the compositions of the feed and the overheadeffluent from the distillation column operating under the conditionsdescribed above.

                  TABLE 7                                                         ______________________________________                                                    Feed    Head                                                        (weight %) (weight %)                                                       ______________________________________                                        <C.sub.4       0.57      0.57                                                   iC.sub.4 14.55 14.66                                                          iC.sub.4.sup.= 84.07 84.69                                                    C.sub.4 .sup.= 1  0.03  0.03                                                  C.sub.4.sup.== 1,3 -- --                                                      nC.sub.4  0.22  0.01                                                          C.sub.4.sup.= 2trans  0.38  0.04                                              Neo C.sub.5 -- --                                                             Me cyclo C.sub.3 -- --                                                        C.sub.4.sup.= 2cis  0.18 --                                                   >C.sub.4 -- --                                                              ______________________________________                                    

These successive and discontinuous hydroisomerisation and distillationoperations represent the separation of 1-butene from isobutene which iscarried out continuously in the process of the invention.

Example 4

Pilot hydroisomerisation tests were carried out using a 1-raffinateusing the hydroisomerisation catalyst LD267R from PROCATALYSE whichpacked each of the catalytic beds. The results of these tests are shownin Table 8 below: they allowed computation parameters to be determinedwhich allowed the process of the invention to be simulated usingsuitable software. The software used for this simulation is sold bySIMCI under the trade name Pro2.

                                      TABLE 8                                     __________________________________________________________________________    pilot test results                                                            __________________________________________________________________________    T ° C.                                                                          40 80 90 50 50 50 50 50 50 50                                          HSV h.sup.-1  30   30   30   30   30   30   30   30   20   40                                                    P bar  10   10   10   6.5 10   15                                            10   10   10   10                           H.sub.2 /HC   0.17  0.17  0.17  0.17  0.17  0.17 0.1  0.19  0.17  0.17                                           m/m                                         feed effl effl effl effl effl effl effl effl effl effl                       <C4  0.14  0.11  0.12  0.11  0.10  0.10  0.10  0.10  0.11  0.10  0.09                                            iC4  5.69  5.75  5.75  5.73  5.71                                            5.76  5.75  5.72  5.76  5.75  5.74                                             iC4= 78.67 78.71 78.72 78.73 78.78                                           78.72 78.73 78.74 78.72 78.71 78.74                                            1-iC4=  3.66  1.30  0.91  0.75  1.15                                          1.01  1.18  1.13  1.00 0.8  1.32                                              n-C4  7.16  7.19  7.17  7.14  7.14                                           7.18  7.19  7.16  7.20  7.19  7.18                                             tr2-  4.36  5.40  5.48  5.40  5.46                                           5.49  5.41  5.46  5.48  5.59  5.37                                             C4=                                        cs2-  0.32  1.54  1.85  2.14  1.66  1.74  1.64  1.69  1.75  1.86  1.56                                           C4=                                      __________________________________________________________________________     where effl = effluent.                                                   

The catalytic hydroisomerising distillation zone comprised 2 or 3catalytic distillation doublets, each of the doublets being of the typeshown in the figure, each doublet being itself constituted by acatalytic cell over which three perforated plates were mounted.

Two examples which were simulated using the calculation were carriedout. They are described below.

Example 4A

The configuration of the unit, comprising three catalytichydroisomerisation beds located inside the column, termed reactiveplates, was as follows:

column with 130 theoretical plates, numbered from top to bottom;

supply to plate no. 90;

the reactive plates were plates 10, 25 and 39. Each contained 7.5 m³ ofcatalyst.

Operating conditions:

Flow rate of liquid supply to column: 292.9 kmole/h;

Reflux ratio: 12;

Column head pressure: 6.2 bars absolute;

Column bottom pressure: 7 bars absolute;

Temperature of supply to column: 59° C.;

Column head temperature: 52° C.;

Column bottom temperature: 64.5° C.;

Temperature of reactive plate no. 10: 53° C.;

Pressure of reactive plate no. 10: 6.6 bars absolute;

Flow rate of liquid traversing reactive plate no. 10: 1660 kmol/h;

Temperature of reactive plate no. 25: 54° C.;

Pressure of reactive plate no. 25: 6.6 bars absolute;

Flow rate of liquid traversing reactive plate no. 25: 1660 kmole/h;

Temperature of reactive plate no. 39: 54° C.;

Pressure of reactive plate no. 39: 6.7 bars absolute;

Flow rate of liquid traversing reactive plate no. 39: 1660 kmole/h.

With this configuration and under those operating conditions, thesimulation produced the following results:

    ______________________________________                                               Column supply                                                                            Column head                                                                             Column bottom                                       (kmole/h) (kmole/h) (kmole/h)                                               ______________________________________                                        <C4      1.12         1.12      0.00                                            iC4 4.46 5.58 0.00                                                            iC4= 110.08 108.07 0.89                                                       C4 = 1 7.53 0.02 0.17                                                         nC4 55.27 0.13 55.23                                                          C4 = 2tr 79.76 0.03 84.56                                                     C4 = 2cis 33.49 0.00 35.91                                                    H.sub.2 1.21 0.00 0.00                                                        Total 292.92 114.95 176.76                                                  ______________________________________                                    

Yield of isobutene at column head: 98.2%

1-butene/isobutene molar ratio at column head: 1.85×10⁻⁴.

Example 4B

The configuration of the unit, comprising two catalytichydroisomerisation beds located inside the column, termed reactiveplates, was as follows:

column with 130 theoretical plates, numbered from top to bottom;

supply to plate no. 90;

the reactive plates were plates 10 and 39. Each contained 7.5 m³ ofcatalyst.

Operating conditions:

Flow rate of liquid supply to column: 292.9 kmole/h;

Reflux ratio: 12;

Column head pressure: 6.2 bars absolute;

Column bottom pressure: 7 bars absolute;

Temperature of supply to column: 59° C.;

Column head temperature: 52° C.;

Column bottom temperature: 64.5° C.;

Temperature of reactive plate no. 10: 53° C.;

Pressure of reactive plate no. 10: 6.6 bars absolute;

Flow rate of liquid traversing reactive plate no. 10: 1660 kmol/h;

Temperature of reactive plate no. 39: 54° C.;

Pressure of reactive plate no. 39: 6.7 bars absolute;

Flow rate of liquid traversing reactive plate no. 39: 1660 kmole/h.

With this configuration and under those operating conditions, thesimulation produced the following results:

    ______________________________________                                               Column supply                                                                            Column head                                                                             Column bottom                                       (kmole/h) (kmole/h) (kmole/h)                                               ______________________________________                                        <C4      1.12         1.12      0.00                                            iC4 4.46 5.17 0.00                                                            iC4= 110.08 108.48 0.89                                                       C4 = 1 7.53 0.09 0.17                                                         nC4 55.27 0.13 55.20                                                          C4 = 2tr 79.76 0.06 84.50                                                     C4 = 2cis 33.49 0.00 35.90                                                    H.sub.2 1.21 0.00 0.00                                                        Total 292.92 115.49 176.66                                                  ______________________________________                                    

Yield of isobutene at column head: 98.6%

1-butene/isobutene molar ratio at column head: 8.30×10⁻⁴.

The preceding examples can be repeated with similar success bysubstituting the generically or specifically described reactants and/oroperating conditions of this invention for those used in the precedingexamples.

The entire disclosures of all applications, patents and publications,cited above and below, and of corresponding application French No.95/15.530, filed Dec. 27, 1995, are hereby incorporated by reference.

From the foregoing description, one skilled in the art can easilyascertain the essential characteristics of this invention, and withoutdeparting from the spirit and scope thereof, can make various changesand modifications of the invention to adapt it to various usages andconditions.

We claim:
 1. A process for the treatment of a feed comprising a majorportion of olefinic hydrocarbons containing 4 carbon atoms per molecule,including isobutene, 1-butene and 2-butenes in a ratio whichsubstantially corresponds to the thermodynamic equilibrium, in whichsaid feed is treated in a distillation zone comprising a stripping zoneand a rectification zone associated with at least one hydroisomerisationreaction zone, said hydroisomerisation reaction zone being at leastpartially internal to said distillation zone and comprising at least onecatalytic bed, in which hydroisomerisation of at least a portion of1-butene is carried out in the presence of a hydroisomerisation catalystand at least one gas stream comprising hydrogen, such that an effluentrich in isobutene leaves the distillation zone overhead and an effluentrich in 2-butenes leaves the bottom, said process being characterized inthat each catalytic bed in the internal portion of thehydroisomerisation zone is traversed by an ascending co-current of saidgas stream and liquid and is substantially out of contact with thedistillation vapor.
 2. A process according to claim 1 in which, for eachcatalytic bed in the internal portion of the reaction zone, the liquidis distributed at a level located below the catalytic bed and the gasstream is distributed at a level below or in the catalytic bed.
 3. Aprocess according to claim 2, in which gas is introduced into the liquidupstream of the level where the liquid is distributed with respect tothe direction of liquid circulation.
 4. A process according to claim 1,in which the gas stream is introduced at a level located substantiallyat the level of where the liquid is distributed, the gas and liquidbeing introduced separately into the catalytic bed.
 5. A processaccording to claim 4, in which the gas stream is introduced in thecatalytic bed.
 6. A process according to claim 4, in which the gasstream is introduced below the catalytic bed.
 7. A process according toclaim 5, in which the gas stream is introduced at a level locatedproximate to the level where the liquid is distributed.
 8. A processaccording to claim 1, in which the hydroisomerisation zone is completelyinternal to the distillation zone.
 9. A process according to claim 1, inwhich the internal portion of the hydroisomerisation zone is at leastpartially in the rectification zone.
 10. A process according to claim 1,in which a major portion of the gas stream is hydrogen.
 11. A processaccording to claim 1, wherein a major portion of 1-butene is subjectedto hydroisomerisation.
 12. A process according to claim 6, in which thegas stream is introduced at a level located proximate to the level wherethe liquid is distributed.